Two stage cracking of residuals



Aug. 17, 1965 A. o, ANDERSON ETAL TWO STAGE CRACKING 0F RESIDUALS Filed NOV. 21. 1960 u f a m m wm v M v mm W V O H lo 9 m A m w m w L? ND Mm HOLVHBNSSHB HOLQVEB m w H T WN LW O O E WSTS n R85 L A MOU FR A.

mm mE VWWf MRHM/ Y B KMCHUIMKQ United States Patent TWO STAGE CRACKING 0F RESIDUALS Arvin D. Anderson, Anaheim, Calif., Robert L. Foster, Homewood, Ill., and Howard G. Russell, Munster, Ind, assignors, by mesne assignments, to Sinclair Research,

Inc., New York, N.Y., a corporation of Delaware Filed Nov. 21, 1960, Ser. No. 70,641 Claims. (Cl. 208-74) This invention is an improved procedure for catalytic cracking of hydrocarbon charge stocks which contain metal impurities. Large quantities of mineral oil petroleum crudes, fractions thereof, and hydrocarbons derived therefrom, contain harmful amounts of metal impurities, such as nickel and vanadium. These impurities are frequently present in such large amounts that utilization of the hydrocarbon is a real problem since the metals accumulate on catalysts in cracking, adversely affecting the product distribution of cracking yields by increasing coke and gas make and decreasing gasoline make. Thus catalytic cracking of the hydrocarbons is uneconomical because the metal impurities harmfully affect selectivity of the catalyst. For this reason, such stocks have not heretofore been utilized to the fullest possible extent.

This invention relates to a multi-stage process for catalytically converting heavy hydrocarbon oils into lower boiling products such as gasoline. More particularly it relates to a process wherein a residual petroleum oil is catalytically cracked in a zone of low severity and later in a zone of higher severity. The residual feedstock employed contains a high proportion of metal contaminants and substantially all of these are deposited on the catalyst in the initial cracking stage. The catalyst employed in the first stage reactor is kept reasonably low in metal poisons by demetallization, that is, by removing a significant amount of poisoning metal from the catalyst. A hydrocarbon fraction from the first cracking stage boiling primarily above about 400 F. and having a relatively insignificant content of poisoning metals is then cracked in a second catalytic cracking unit, where a high severity can be maintained for maximum gasoline production without encountering a prohibitive metal contamination problem.

The need has been expressed in the art for a feasible process for the catalytic Cracking of petroleum residua or similar heavy hydrocarbon feeds. The chief deterrent to catalytic cracking of residua by conventional means has been the severe catalyst contamination, due to cokeformers and contaminant metals in most residua, which leads to poor catalyst activity and often to poor product distribution and otherwise reduces the desired effectiveness of the catalyst. Consequently, two-stage vapor phase catalytic processes have been proposed according to which the coke is supposed to be deposited on a catalyst in a first stage to minimize the adverse eflfect of coke formation in the second stage, and the resulting purified feed is only thereafter subjected to this second catalytic cracking wherein the main conversion is supposed to be accomplished. However, the difficulty with this type of operation has been that there was no successful process for carrying out the first catalytic cracking step at sufficiently low conversion and without undue loss of gas oil components of the product through overcracking, dehydrogenation, etc.; with the result that the feed conversion in the first stage approached the second stage conversion in severity and the production of suitable feeds for the second stage was ineflicient. The diificulty lay largely in the fact that in order to achieve the complete vaporization necessary for a vapor phase catalytic cracking process, such as a conventional fluid catalytic cracking, high temperatures, high conversion of the residua, and high metal contamination of catalyst was unavoidable.

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More recent attempts to Work out a practical process for converting heavy residua have all but abandoned the idea of direct catalytic cracking. Instead, the trend has been toward first coking the residua in the presence of inert solids and subjecting only the liberated distillate to a catalytic conversion step. Of course, this also has meant considerably poorer product distribution and poorer prod uct quality than in an all-catalytic conversion; especially since the inert solids become poisoned with metal contaminants and therefore exert some unwanted catalytic activity.

Thus, it is the main object of this invention to provide a practical, two-stage catalytic process for converting heavy hydrocarbons to gasolines and other desired products. A more specific object is the devising of a multistage process wherein the hydrocarbon feed is catalytically cracked in the mixed state, i.e. a mixed liquid-vapor phase, at low conversions in a first stage; the resulting product; substantially free of metal poisons, is separated from the first stage catalyst with minimum chemical change, and

I its heavier portion is subsequently cracked in the presence of an active catalyst; and the first stage catalyst is demetallized.

The present invention obviates many of the difliculties which have previously plagued the oil refiner; metal contamination of the catalyst is substantially confined to the first stage, therefore only part of the catalyst inventory needs treatment for demetallization; the severity of operation for the two stages can be balanced for maximum yield of desired products. Coke yield may be held to a minimum through use of good steam stripping and a high steam partial pressure in the initial reactor, and the second stage, where the main cracking is carried out, is substantially free of a metal poisoning problem.

One of the most important phases of the study in the improvement of catalyst performance in hydrocarbon conversion is in the area of metals poisoning. Although referred to as metals, these catalyst contaminants may be in the form of free metals or relatively non-volatile metal compounds, It is to be understood that the term metal used herein refers to either form. Various petroleum stocks have been known to contain at least traces of many metals. For example, Middle Eastern crudes contain relatively high amounts of several metal components, while Venezuelan crudes are noteworthy for their vanadium content and are relatively low in other contaminating metals such as nickel. Because most of these metals, when present in a stock, would deposit'in a relatively nonvolatile form on the catalyst during the conversion processes and regeneration of the catalyst to remove coke would not remove these contaminants, such feeds are generally avoided. Nickel and vanadium for example, markedly alter the selectivity and activity of cracking reactions if allowed to accumulate, producing a higher yield of coke and hydrogen at the expense of desired products, such as gasoline and butanes.

When an attempt is made to segregate higher boiling distillate fractions of a crude oil, it is found that the metal contaminants tend to concentrate in the heavier fraction or residuum. A minor portion of metal contaminants is inherently and unavoidably carried over into the distillate products, but such distillate fractions are frequently used as cracking feedstocks, with the catalyst poisoning problem being solved by high rates of catalyst replacement. Contaminated residual fractions are generally used as sources of asphalt, fuel, and other products which are of relatively low economic value and where the presence of metal contaminants is not unduly harmful. Contaminated crudes or other residuals therefore are not usually fully exploited as cracking feeds. This invention, however, provides for the upgrading of Patented Aug; 17, 1965 C.) such residual material to gasoline and other desirable products by an economically feasible process.

Solid oxide catalysts, both naturally occurring activated clays and synthetically prepared gel catalysts, as well as mixtures of the two types, have long been recognized as useful in catalytically promoting conversion of hydrocarbons. A popular natural catalyst is Filt-rol which is acid-activated montmorillonite. For cracking processes, the solid oxide catalysts which have received the widest acceptance today are usually activated or calcined predominantly silica or silica-based, e.g. silica-alumina, silicia-magnesia, silicia-zirconia, etc., compositions in a state of slight hydration and containing small amounts of acidic oxide promoters in many instances. The oxide catalyst may be alumin-aor silica-based and ordinarily contains a substantial amount of a gel or gelatinous precipitate comprising a major portion of silica and at least one another material, such as alumina, zirconia, etc. The use of wholly or partially synthetic gel catalysts, which are more uniform and less damaged by high temperatures in treatment and regeneration, is often preferable. Popular synthetic gel cracking catalysts generally contain about to 30% alumina. Two such cataly-asts are Aerocat which contains about 13% A1 0 and High Alumina Nalcat which contains about 25% A1 0 with substantially the balance being silica. The catalyst may be made only partially of synthetic material; for example it may be made by the precipitation of silica-alumina on clay, such as k-aolinite or halloysite. One such semisynthetic catalyst contains about equal amounts of silicaalumina gel and clay.

The physical form of the catalyst varies with the type of manipulative process to which it will be exposed. In fluid catalytic processes gases are used to convey the catalyst in the form of a fine powder, generally in a size range of about to 150 microns. In the TCC or Thermofor process the catalyst is in the form of beads which are conveyed by elevators. Generally these beads may range in size up to about /2" in diameter. When fresh, the minimum sized bead is generally about /s". Other types of process use other forms of catalyst such as tablets or extruded pellets. The cracking process of this invention preferably uses the fluidized solids technique.

The feeds to the present process comprise residua which may be exemplified by vacuum residua, atmospheric residua, tars, pitches, etc., boiling essentially above about 400 or even above about 600 F., and will in general contain greater amounts of poisoning metals than in distillate oils. The feed often has an API gravity in the range of about 0 to a Conradson carbon content in the range of about 3 to weight percent and a viscosity above about 75 or even 200 seconds Saybolt Furol at 210 F. These charge stocks, containing more than about 1.5 parts per million of vanadium and/or more than 'about 0.6 part per million of nickel are generally avoided in cracking processes and most refiners prefer less than about 0.5 part per million of vanadium or about 0.2 part per million of nickel in the cracking stock. Metal contents above these ranges may be present in feeds to the process of this invention; it will be apparent that oils having metal contents in these generally undesirable ranges are the oils which this invention salvages. A mixture of vanadium and nickel may be considered as harmful as a single metal even though the individual amounts of each metal are below the values mentioned above because the effect of the total amount of the metallic components is frequently sufficient to give harmful effects during catalytic cracking. In most cases, however, the total of one, two or more of these metals in the residual will be at least about 0.5 p.p.m. The maximum amount of metals in the residuals can vary widely; most often the maximum amount of these poisoning metals in the residual stock will not exceed about 50 p.p.m. nickel, and about 100 vanadium.

In the first stage of cracking, diluents for the feed may be employed. Such diluents include steam, other inert gases, and low end-point hydrocarbon vapors and may be as much as about 150 weight percent of the petroleum residual and when used are usually at least about 10 percent. Such vapors frequently are low octane naphthas or light gas oils or other hydrocarbons which are normally liquid but are readily volatile since primarily they do not exceed the gas oil boiling range. When naphthas are added to the hydrocarbon feed, the amount may vary from about 0 to 60 gal, preferably 5 to 20 gal., of naphtha per barrel of residuum feed. The use of such vapors as diluents permits a more complete cracking of the heavior stock with less coke formation and promotes the vapor ization of more of the heavier material in the feedstocks. Steam may be added in amounts ranging from about 0 to 50 weight percent, preferably about 20 to 40 weight percent, based on the amount of hydrocarbon feed. In the practice of this invnetion, a mixture of hydrocarbon feed and catalyst, with about to lbs. of steam per barrel of residuum can be passed to the first cracking zone.

In the first cracking zone the feedstock is subjected to mild cracking conditions sufficient to lay down on the cracking catalyst substantially all of the metals in the feed, that is, metals are removed from the residual feed to such an extent that the feed to the second stage cracking zone contains no more poisoning metals than conventional gas oil feeds boiling in the range of about 400 to 1100l200 F., for instance, about 0.4 p.p.m. nickel and 0.8 p.p.m. vanadium and preferably about 0.2 p.p.m. nickel, and 0.4 p.p.m. vanadium. The associated coke also is generally deposited on the catalyst in this mild cracking stage. The actual amount of feedstock cracking which occurs in this first stage varies widely with the crude source and boiling range of the feedstock. However, the extent of cracking in this first stage need be no greater than about 30-45% for adequate removal of contaminants from the feed, that is no more than about 3045% of the feed need be converted to products boiling below about 400 F. in order to decontaminate the residual oil feedstock. This invention provides for control of the extent of cracking in the first cracking stage by employing one or more specialized techniques. Temperature is controlled so that the feedstock is contacted with catalyst in the cracking zone at a temperature between about 700 to 900 F., preferably about 800 to 875 F. The time during which feed and catalyst are in contact under cracking conditions varies; a high reactor space velocity of about 1020 WHSV may be useful to help control the severity of the cracking although a lower space velocity may be used where the catalyst is less active. Other means of holding conversion to a minimum in the first stage may include, besides controlling the temperature and space velocity, for instance, the injection of the catalyst at the top of the reactor riser at a point just be fore entering the reaction while conveying the catalyst with an inert gas such as steam to minimize contact of catalyst and feed. The actual conversion obtainable will vary widely with the crude source and its boiling range, but at least one of the above-mentioned means may be used to keep conversion at a desirable minimum, for example, a conversion of the residual to lighter materials of up to about 30-45%. The cracking is generally conducted at pressures of about 5-15 p.s.i.g. and the feed may be heated to a temperature of, say, about 600 to 700 F. before introduction to the cracking zone. This temperature is generally not sufficient to vaporize all of the feedstock. The cracking is generally conducted under fluidized catalyst conditions while continuously removing from the first stage cracker a vaporous effluent which contains no significant amount of poisoning metals.

Preferably, coke yield is held to a minimum through the use of good steam stripping and a high steam partial pressure in the reactor and the catalyst surface is kept reasonably free of metal poisons by demetallization. Ordinarily, the catalysts are treated before the poisoning metals have reached an undesirably high level, for instance, a-bout 2%, generally no more than about 1% maximum, content of vanadium and/or nickel.

The catalyst may be removed from the first stage cracker-that is, the stream of catalyst which is cycled between conversion .and regenerating operations-before the poison content reaches about 5000 to 10,000 p.p.m., the poisoning metals being calculated as their common oxides. Generally, at least about 250 or 500 ppm. vanadium and/or at least about 100 or 200 p.p.m. nickel will be accumulated on the catalyst before demetallization is warranted. A suitable amount, generally a small portion of the catalyst, is removed from the first stage cracker and demetallized, preferably after the oxidation regeneration which serves to remove carbonaceous deposits. With a continuously circulating catalyst stream, such as in the ordinary fluid system a slip-stream of catalyst may conveniently be removed intermittently or continuously from the regenerator standpipe. The severity of regeneration is generally such that the catalyst sent to demetallization contains not more than about 0.5% carbon. Where the catalyst is sent to a demetallization which includes a high temperature treatment with molecular oxygen-containing gas before the catalyst is substantially carbon free, the length of oxygen treatment is reckoned from the time that the catalyst reaches the substantially carbon free state, that is the state where little, if any, carbon is burned or oxygen consumed even when the catalyst is contacted with oxygen at temperatures conducive to combustion.

In the practice of this invention demetallization may be accomplished by the intermittent or continuous withdrawal of contaminated catalyst from the initial cracking sytsem, for example, from the regenerator standpipe. The catalyst is subjected to one or more of the demetallization procedures described hereinafter, and then the catalyst, substantially reduced in contaminating metal content is returned to the first stage cracking system. In the treatment to take poisoning metals from the cracking catalyst a large or small amount of metal can be removed as desire-d. The amount of poisoning metal that is removed from the catalyst depends on the amount of poison which can be tolerated, for example, where a cracking unit can tolerate 100 p.p.m. Ni and the demetallization process removes 50% of the metals, only 50 p.p.m. Ni is removed in each demetallization, but where a unit can tolerate 500 p.p.-m., 250 ppm. may be removed each time the poisoned catalyst is sent through the demetallization unit. The tolerance of the unit for poison is determined, for example, by the loss in yield due to poisoning compared to the cost of enlarged demetallization facilities. When treating the catalyst it is preferred to remove at least about up to about 70% or more of one or more of the metals in question. The demetallization rate may be about 20-200% of catalyst inventory daily, that is, within a 24-hour period, more than about /s of the amount of catalyst in the system is subjected to the demetallization and sometimes, when it is necessary, twice the amount of catalyst in the system is subjected to one or more of the demetallization procedures.

The hydrocarbon products from the first stage cracking zone are conducted to a fractionator, with or without intermediate cooling. In the fractionation, which may be conducted under partial vacuum, the lower boiling gasoline constituents of the cracker efiluent having an approximate 375 -430 F. end point are vaporized and removed from the system and may be used as gasoline blending components or other products. The liquid hydrocarbon that is charged to the principal cracking zone and substantially metals free may be a refractory cycle stock usually boiling between about 400 F. and about 850-950 F. or the feed may be the entire portion of the residual remaining after the light ends (the portion boiling up to about 400 F.) have been removed; therefore, such a fraction may include the entire gas oil fraction (400 F. to 1100- 1200 F.) and the undistilled portion of the residual boiling above about 1100-l200 F. The feed to the principal cracking zone is vaporized and catalytically treated under more or less conventional fluid catalytic cracking conditions. These conditions generally include a temperature at least about 25 higher than the temperature main tained in the first stage, at about 850 to 1000 F., preferably about 925 to 975 F. and a pressure between about 5 and 25 p.s.i.g., at a weight hourly space velocity from about 0.1 to 10 to obtain about a -80 volume percent conversion of the 400 F. gas oil to gasoline and other desired lighter components. The products of the second stage cracking are conducted to a fractionator. The gasoline fraction may be joined with the 375430 F. end point cut obtained from the first cracking and the heavier products may be recycled to the preliminary or main cracking zone for further processing.

The accompanying drawing illustrates in schematic form apparatus which may be used in performing this invention. Residual feed is conducted by line 10 to the preheater 12 where the temperature of the feed may be raised almost to cracking temperature as described above. Heated feed, generally as a mixture of liquids and vapors is conducted by line 14 to the cracking reactor 16 where the preliminary mild catalytic cracking is performed.

Catalyst, when a continuous cracking system is employed,

is continuously sent by line 13 to regenerator 20 where carbon is burned, for example by contact with air, from the catalyst. Catalyst returns to the reactor by line 22.

Line 24 is provided for removal of a side stream of catalyst from the line 22. This catalyst is sent to a de-- metallization unit 26. This unit represents one or a plurality of treating zones wherein poisoning metal is prepared for removal or removed from the catalyst by one or more of the demetallization procedures described below. The catalyst, of reduced metal content, returns to the cracking system, for example to the regenerator 20,

a by means of the line 28.

Line 30 conducts effluent vapors from the first stage cracking to the fractionator 32 which provides for removal of lower boiling constituents of the cracker effiuent by line 34. The fraction boiling above the gasoline range travels by line 36 to the principal cracking reactor 38. A portion of the cracking catalyst may be continuously or intermittently removed from reactor 38 by line 44 to regenerator 42 and held, for example, in air at a temperature and for a sufficient amount of time to substantially reduce the carbon content of the catalyst and then returned to reactor 38 by line 40.

Eiiiuent gases from the principal cracker are conducted by line 46 to the fractionator 48, from which heavier products may be recycled by line 50 to the preliminary cracker while gasoline and lighter fractions are removed by line 52.

Regeneration of a catalyst to remove carbon is a relatively quick procedure in most commercial catalytic con version operations. For example, in a typical fluidized cracking unit, a portion of catalyst is continually being removed from the reactor and sent to the regenerator for contact with air at about 950 to 1200 F., more usually about 1000 to 1150 F. Combustion of coke from the catalyst is rapid, and for reasons of economy only enough air is used to supply the needed oxygen. Average residence time for a portion of catalyst in the regenerator may be on the order of about six minutes and the oxygen content of the efliuent gases from the regenerator is desirably less than about /2%.

As mentioned, a minor portion of the catalyst is continuously removed as a side stream from the regenerator associated with the first stage cracker and is sent to a demetallization system. A number of procedures have become available which provide for the removal of the common metal poisons iron, nickel and vanadium from a contaminated hydrocarbon conversion catalyst. Such procedures frequently involve contact of the catalyst at elevated temperatures with vapors reactive with the poisoning metal on the catalyst and are described, for example, in copending applications Serial Nos. 763,834, filed September 29, 1958, now abandoned; 767,794, filed October 17, 1958, now abandoned; 849,199, filed October 28, 1959, now abandoned; 842,618, filed September 28, 1959, now abandoned; 19,313, filed April 1, 1960, now abandoned; 39,810, filed June 30, 1960; 47,598 filed August 4, 1960; 53,623, filed September 2, 1960; 55,129, filed September 12, 1960, now Patent No. 3,147,209; 55,160, filed September 12, 1960; 55,380, filed September 1, 1960, now Patent No. 3,122,497; 54,532, filed September 7, 1960, now abandoned; 55,838, filed September 14, 1960, now abandoned; all of which are herein incorporated by reference. It has been found, for example, that Ni and V may be removed from a catalyst by converting the metals into volatile compounds;.a chlorination treatment can convert the vanadium to volatile chlorides, as reported in copending applications Serial Nos. 849,199, filed October 28, 1959, and 54,532, filed September 7, 1960, and nickel may be converted to the volatile nickel carbonyl by hydrogenation and treatment with carbon monoxide, as pointed out in copending application Serial No. 47,598, filed August 4-, 1960.

It has also been found that Ni and V may be removed from a catalyst by certain aqueous media; a basic aqueous wash containing ammonium ions is suitable for removal of V poisons as reported in copending application Serial No. 39,810, filed June 30, 1960. The removal of nickel may be accomplished by the use of a slightly acid aqueous wash when the nickel is first converted into a compound dispersible in such a wash. The chlorination treatment can convert nickel to the soluble chloride form. Also, a poisoned catalyst may be reduced in nickel content by the aqueous wash when nickel contaminants are put into the sulfate or other dispersible form by oxidizing a sulfided nickel-contaminated catalyst. Such an oxidation may be performed by an oxidizing vapor, as disclosed in copending applications Serial Nos. 763,834, filed September 29, 1958, and 55,129, filed September 12, 1960, or by an aqueous oxidizing agent, as explained in copending application Serial No. 842,618, filed September 28, 1959. Sulfidation of nickel poisoned catalyst appears to have important effects in making more nickel available for removal, so that sulfidation, as described in the latter two copending applications and in application Serial No. 53,330, filed September 1, 1960, may be performed when the nickel removal is by a route other than conversion to sulfate. Also, as pointed out in copending applications Serial No. 19,313, filed April 1, 1960, and Serial No. 55,160, filed September 12, 1960, a preliminary treatment of the catalyst with molecular oxygen-containing gas is of value in improving the vanadium removed by subsequent procedures. The treatment may remove a substantial part of one or both of these metals. The withdrawal of catalyst from the cracking system can be on a continuous or intermittent basis and ordinarily the catalyst will not be allowed to accumulate more than about 5000 or 7500 ppm. of poisoning metal. Subjecting the poisoned catalyst sample to magnetic flux may be found desirable to remove any tramp iron particles which may have become mixed with the catalyst.

Treatment of the regenerated catalyst with molecular oxygen-containing gas is described in copending applications Serial Nos. 19,313, filed April 1, 1960, and 55,160, filed September 12, 1960, hereby incorporated by reference. The temperature of this treatment is generally in the range of about 1000 to 1800 F. but below a temperature where the catalyst undergoes any substantial deleterious change in its physical or chemical characteristics. The catalyst is in a substantially carbon-free condition during this high-temperature treatment and there is substantially no oxygen consumption. If any significant amount of carbon is present in the catalyst at the start of this high-temperature treatment as Where the extent of regeneration is limited to decrease catalyst activity in the first cracking stage, the essential oxygen contact is that continued'after carbon removal. In any event, after carbon removal, the oxygen treatment of the essentially carbon-free catalyst is at least long enough to convert a substantial amount of vanadium to a higher valence state, as evidenced by a significant increase, say at least about 1 96, preferably at least about in the vanadium removal in subsequent stages of the process. This increase is over and above that which would have been obtained by the other metals removal steps without the oxygen treatment.

The treatment of the vanadium-poisoned catalyst with molecular oxygen-containing gas is preferably performed at a temperature of about 1150 to 1350 or even as high as 1600 F. and at least about 50 F. higher than the regeneration temperature. The oxygen-containing gas used in the treatment contains molecular oxygen as the essential active ingredient. The gas may be oxygen, or a mixture of oxygen with inert gas, such as air or oxygen-enriched air. The partial pressure of oxygen in the treating gas may range widely, for example, from about 0.1 to 30 atmospheres, but usually the total gas pressure will not exceed about 25 atmospheres. The factors of time, partial pressure and extent of vanadium conversion may be chosen with a view to the most economically feasible set of conditions. It is preferred to continue the oxygen treatment for at least about 15 or 30 minutes with a gas containing at least about 1%, preferably at least about 10% oxygen. The catalyst may pass directly from the oxygen treatment to an ammonia wash as described below for satisfactory vanadium removal especially where this is the only important contaminant, as may be the case when the residum cracked is derived, for example, from Venezuelan crude.

Conversion of the metal poisons to the sulfate or other dispersible form is described in copending applications Serial No. 763,834, filed September 29, 1958, and Serial No. 842,618, filed September 28, 1959, hereby incorporated by reference, and may be accomplished, for instance, by subjectin the catalyst to a sulfating gas, that is, 50 S0 or a mixture of S0 and 0 at an elevated temperature. Sulfur oxide contact is usually performed at a temperature of about 500 to 1200 F. and frequently it is advantageous to include some free oxygen in the treating gas. Another procedure includes sulfiding the catalyst and converting the sulfide by an oxidation process, after which metal contaminants in dispersible form may be removed from the catalyst by an aqueous medium.

A sulfiding can be performed by contacting the poisoned catalyst with elemental sulfur vapors, or more conveniently by contacting the poisoned catalyst with a volatile sulfide, such as H 5, CS or a mercaptan. The contact with the sulfur-containing vapor can be performed at an elevated temperature generally in the range of about 500 to 1500 F, preferably about 800 to 1300 F. Other treating conditions can include a sulfur-containing vapor partial pressure of about 0.1 to 30 atmospheres or more, preferably about 0.5 to 25 atmospheres. Hydrogen sultide is the preferred sulfiding agent. Pressures below atmospheric can be obtained either by using a partial vacuum or by diluting the vapor with gas such as nitrogen or hydrogen. The time of contact may vary on the basis of the temperature and pressure chosen and other factors such as the amount of metal to be removed. The sulfiding may run for, say, up to about 20 hours or more depending on these conditions and the severity of the poisoning. Temperatures of about 900 to 1200 F. and pressures approximating 1 atmosphere or less seem near optimum for sulfiding and this treatment often continues for at least 1 or 2 hours but the time, of course, can depend upon the manner of contacting the catalyst and sulfiding agent and the nature of the treating systern, e.g. batch or continuous, as well as the rate of diffusion within the catalyst matrix.

The sulfiding step performs the function not only of supplying a sulfur-containing metal compound which may be easily converted to the sulfate or other dispersible form but also apparently serves to concentrate some metal poisons, especially nickel, at the surface of the catalyst particle.

Oxidation after sulfiding may be performed by a gaseous oxidizing agent to convert metal sulfide to sulfate, including oxysulfate, or other dispersible form. Gaseous oxygen, or mixtures of gaseous oxygen with inert gases such as nitrogen, may be brought into contact with the sulfided catalyst at an oxygen partial pressure of about 0.2 atmosphere and upward, temperatures upward of room temperature and usually not above about 1300 F., and times dependent on temperature and oxygen partial pressure. Gaseous oxidation is best carried out near 900 F., about one atmosphere O and at very brief contact times.

The metal sulfide may be converted to the corresponding sulfate, or other dispersible form, by a liquid aqueous oxidizing agent such as a dilute hydrogen peroxide or hypochlorous acid Water solution as described in copending application Serial No. 842,618, filed September 28, 1959. Bromine, chlorine, or iodine water, or aerated, ozonatedor oxygenated water, with or without acid, also will oxidize the sulfides to sulfates or other dispersible form. The inclusion in the liquid aqueous oxidizing solution of sulfuric acid or nitric acid has been found greatly to reduce the consumption of peroxide. In addition the inclusion of nitric acid in the oxidizing solution provides for increased vanadium removal. Useful proportions of acid to peroxide to catalyst generally include about 2 to 25 pounds acid (on a 100% basis) to about 1 to 30 pounds or more H 0 (also on a 100% basis) in a very dilute aqueous solution, to about one ton of catalyst. A 30% H 0 solution in water seems to be an advantageous raw material for preparing the aqueous oxidizing solution. Another highly advantageous oxidizing medium is an aerated dilute nitric acid solution in Water. Such a solution may be provided by continuously bubbling air into a slurry of the catalyst in very dilute nitric acid. Other oxygen-containing gases may be substituted for air. least about 7 to 8 minutes. The oxidation slurry may contain about 20% solids and provide about five pounds of nitric acid per ton of catalyst. The liquid phase oxidation may also be performed by exposing the sulfided catalyst first to air and then to the aqueous nitric acid solution. The conditions of oxidation can be selected as desired. The temperature can conveniently range up to about 220 F. with temperatures of above about 150 F. being preferred. Contact with the hot catalyst may be sufficient to raise the temperature of the slurry from ambient temperature to around the boiling point. Temperatures above about 220 F. necessitate the use of superatmospheric pressures and no need for such has been found.

Sometimes removal of sulfides from the catalyst may be accomplished by contacting the catalyst with an ap propriate solvent. Such procedures are described in copending application Serial No. 763,833, filed September 29, 1958, now abandoned, incorporated herein by reference. These solvents are in general aqueous and may contain a complexing or chelating agent for the nickel and/or other metal poisons. Aqueous solutions containing cyanide or hexametaphosphate ions are useful in forming soluble complexes with the poisoning metals. Organic sequestering agents, such as ethylene diamine tetraacetic acid (EDTA), etc. have been found useful in removing the sulfided metals since they form soluble chelate complexes with the metals and effectively retard redeposition of the poisoning metals on the catalyst surface once they are brought into solution. The liquid The time required for oxidation is generally at 10 phase aqueous medium may be applied to the sulfided catalyst at any temperature from ambient temperature upwards. Elevated temperatures approaching the boiling point of Water are preferred. Also, it has been found desirable sometimes to impart oxidation characteristics to the wash containing a chelating agent.

It has further been found that treatment of a metalscontaminated catalyst With a chlorinating agent at a moderately elevated temperature is of value in removing vanadium and iron contaminants from the catalyst as volatile chlorides. This treatment is described in copending application Serial No. 849,199, filed October 28, 1959. Generally, the major proportion of these volatile chlorides is removed during contact with the chlorinating vapor and where the volatile chlorides are insufiiciently removed, a purge with an inert gas such as nitrogen at an elevated temperature may be applied to the chlorinated catalyst. A conversion to vanadium chloride after the high temperature oxygen and/ or sulfiding treatment pref erably makes use of vapor phase chlorination at a moderately elevated temperature wherein the catalyst composition and structure is not materially harmed by the treatment and a substantial amount of the poisoning metals content is converted to chlorides. The chlorination takes place at a temperature of at least about 300 F. to say about 1000 F., preferably about 550 to 650 F., with optimum results being obtained close to about 600 F. The chlorinating reagent is a vapor which contains chlorine, preferably in combination with carbon or sulfur. Such reagents include molecular chlorine but preferably are the chlorine substituted light hydrocarbons, such as carbon tetrachloride, which may be used as such or formed in-situ by the use of, for example, a vaporous mixture of chlorine gas with low molecular weight hydrocarbons such asmethane, ethane and propane. The chlorination may take about 5 to 120 minutes, more usually about 20 to 60 minutes, but shorter or longer reaction periods may be possible or needed, for instance, depending on the linear velocity of the chlorinating and purging vapors.

Nickel poison may be removed by conversion of the nickel sulfide to the volatile nickel carbonyl by treatment with carbon monoxide, as described in copending applica-' tion Serial No. 47,598, filed August 4, 1960, incorporated herein by reference. treated with hydrogen at an elevated temperature during which nickel contaminant is reduced to the elemental state, then treated, preferably under elevated pressure and at a lower temperature, with carbon monoxide, during'which nickel carbonyl is formed and flushed off the catalyst surface. Some iron contaminent is also removed by this carbonylation treatment. 7

' Hydrogenation takes place at a temperature of about 800 to 1600" R, at a pressure from atmospheric or less up to about 1000 p.s.i.g. with a vapor containing 10 to 100% hydrogen. Preferred conditions are a pressure up to about 15 p.s.i.g. and a temperature of about 1100 to 1300 F. and a hydrogen content greater than about mole percent. The hydrogenation is continued until surface accumulations of poisoning metals, particularly nickel, are substantially reduced to the elemental state.

Carbonylation takes place at a temperature substantially lower thanthe hydrogenation, from about ambient temperature to 300 F. maximum and at a pressure up to about 2000 p.s.i.g., with a gas containing about 50 100 mole percent CO. Preferred conditions include greater than about mole percent CO, a pressure of up to about 800 p.s.ig. and a temperature of about F. The CO treatment generally serves to convert the elemental metals, especially nickel and iron, to volatile carbonyls and to remove the carbonyls.

After the conversion of some of the poisoning metal' to a form soluble or dispersible in an aqueous medium, the catalyst can be washed with such aqueous medium to remove metal sulfate, nitrate, etc. or the soluble metal chloride produced in the chlorination procedure In such a procedure the catalyst is described above.

This aqueous wash medium will preferably be somewhat acidic, and this character can be imparted, at least initially, due to the presence of the acid-acting salt or some entrained acidic oxidizing agent on the catalyst. Ambient temperatures can be used in the wash. Pressures above atmospheric may be used but the results usually do not justify the additional equipment. Where an aqueous oxidizing solution is used, the solution may perform part or all of the metal compound removal simultaneously with the oxidation. In order to avoid undue solution of alumina from the catalyst, when the metal poisons have been converted to the chlorides, contact time is preferably held to about 3 to 5 minutes which is sufficient for nickel removal. Also, since a slightly acidic solution is desirable for nickel removal, this wash preferably takes place before an ammonium wash, hereinafter described.

Vanadium may be removed from the catalyst by washing it with a basic aqueous solution as described in copending application Serial No. 39,810, filed June 30, 1960, and incorporated herein by reference. The pH is frequently greater than about 7.5 and the solution preferably contains ammonium ions which may be NH ions or organic-substituted NH ions such as methyl ammonium and quaternary hydrocarbon radical ammoniums. An aqueous solution of ammonium hydroxide is preferred. The preferred solutions have a pH of about 8 to 11. The amount of ammonium ion in the solution is sufficient to give the desired vanadium removal and will often be in the range of about 1 to or more pounds per ton of catalyst treated. Five to fifteen pounds is the preferred ammonium range but the use of more than about 10 pounds does not appear to increase vanadium removal unless it increases pH.

After a wash treatment the catalyst slurry can be filtered to give a cake which may be reslurried with water or rinsed in other ways, such as, for example, by a water wash on the filter, and the rinsing may be repeated, if desired, several times.

After demetallization, the catalyst is conducted back to the first stage cracking, for instance, to the hydrocarbon conversion reactor or catalyst regenerator, although it may be desirable first to dry the catalyst filter cake or filter cake slurry at say about 250 to 450 F. and also, prior to reusing the catalyst in the conversion operation it can be calcined, say at temperatures usually in the range of about 700 to 1100 F. A fluidized solids technique is recommended for the sulfiding and other vapor contact processes used in any selected demetallization procedure as a way to shorten the time requirements. After the available catalytically active poisoning metal has been removed, in any removal procedure, further reaction time may have relatively little effect on the catalytic activity of the depoisoned catalyst, although further metals content may be removed by repeated or other treatments or by increasing the demetallization rate, that is, the fraction of catalyst inventory sent to demetallization per unit time. Inert gases frequently may be employed after contact with reactive vapors to remove any of these vapors entrained in the catalyst or to purge the catalyst of reaction products. Any given step in the demetallization treatment is usually continued for a time sufficient to elfect a substantial conversion or removal of poisoning metal and ultimately results in a substantial increase in metals removal compared with that which would have been removed if the particular step had not been performed. The actual time or extent of treating depending on various factors and is controlled by the operator according to the situation he faces, e.g. the extent of metals content in the feed, the level of conversion unit tolerance for poison, the sensitivity of the particular catalyst toward a particular phase of the demetallization procedure, etc.

The amount of Ni and/or V removed in practicing the procedures outlined or the proportions of each which are removed may be varied by proper choice of treating conditions. it may prove necessary, in the case of very severely poisoned catalysts, to repeat the treatment to reduce the metals to an acceptable level, perhaps with variation where one metal is greatly in excess. A further significant advantage of these demetallization techniques lies in the fact that the overall metals removal operation, even if repeated, usually does not unduly deleteriously affect the activity, selectivity, pore structure and other desirable characteristics of the catalyst.

The present invention will be further described with reference to the following examples which are not to be considered limiting.

Example I A 40% reduced North Texas petroleum crude having an A?! gravity of about 22, a Conradson carbon of about 5 weight percent, a viscosity of about 100 seconds Saybolt Universal at 210 F., and an initial boiling point above about 650 F. at atmospheric pressure, containing 25.0 p.p.m. of nickel and 60 p.p.m. of vanadium is preheated to about 600 to 700 F., and introduced into a first stage cracker, mixed with a fine-1y divided cracking catalyst and about 100 lbs. steam/bbl. residual feed for dispersion and stripping. The catalyst introduced into the feed line is a Nalcat synthetic gel cracking catalyst containing 25% A1 0 the balance silica, and having fluidizable particle size and a relatively low level of activity.

Cracking conditions in the preliminary cracker are low enough to keep conversion of the feedstock to a 430 F. end point gasoline at a 36-42% yield, but sufficiently high to enable the contaminated feedstock to deposit substantially all of its metal contents on the catalyst. The temperature of the first stage zone is kept at about 850 F. under a pressure of about 10 p.s.i.g. The feedstock mixture is conveyed through the cracker at a weight hourly space velocity of about 15. A portion of catalyst is continually sent to a regenerator where it is contacted with air at 1050 F. to burn off the carbon. A side stream of regenerated catalyst having a carbon content of about 0.4%, 375 p.p.m. nickel and 1950 ppm. vanadium is continuously removed from the regenerator at an inventory rate of daily and sent to the demetallization unit. In the demetallization unit the catalyst is first held for about an hour in contact with air at about 1300 F. and then sent to a sulfiding zone where it is fluidized with H S gas at a temperature of about 1100 F. for about an hour. The catalyst is then purged with flue gas at a temperature of about 575 F. and chlorinted in a chlorination zone with an equimolar mixture of C1 and CCl at about 600 F. After about 1 hour no trace of vanadium chloride can be found in the chlorination effluent and the catalyst is quickly washed with water. A pH of about 2 is imparted to this wash medium by chlorine entrained in the catalyst and the wash serves to remove nickel chloride.

The catalyst, substantially reduced in nickel and vanadium content, is filtered from the wash slurry, dried at about 350 F. and returned to the regenerator. The treated catalyst is analyzed and shows a metals content of 151 ppm. nickel and 1420 p.p.m. vanadium.

The vapor products from the preliminary treatment are conducted to a fractionator where the lower boiling constituents from the efiiuent, boiling below about 430 F., are vaporized and removed from the system for use as a gasoline blending component. The total hydrocarbon residue from the fractionator, analyzing about 0.25 ppm. nickel and 0.5 ppm. vanadium and having a boiling point above 430 F. is sent to a second cracking stage where it is treated in the presence of a Nalcat synthetic gel silica-alumina catalyst, having an A1 0 content of about 25%, at a temperature of about 950 to 975 F. and a pressure of about 5 p.s.i.g., at a weight hourly space velocity of about 0.1 to l. The cracked products from the second stage are introduced to a fractionator 13 where a 75% yield of gasoline and other desired components based on the feed to the second stage cracker are removed. The residue may be recycled to the first stage cracker for further processing. A portion of the silica-alumina catalyst is removed from the second cracking reactor to a regenerator and held in a free oxygencontaining gas for about minutes at a temperature of about 1100 F. and then returned to the second stage reactor.

Example II In another run a vacuum residuum was employed, derived from a West Texas crude oil and having an API gravity of 15.1, a Conradson carbon content of about 8.8 weight percent, a viscosity of about 400 seconds Saybolt at 210 F., and an initial boiling point above about 1000 F. and containing 24.7 p.p.m. of nickel and 39.9 p.p.m. of vanadium. A steam-residuum mixture having a 1 to 1 volume ratio is mixed with the cracking catalyst having a particle size range of about 20 to 150 microns. The cracking catalyst comprised a synthetic silica-alumina composite of relatively low activity containing about 13% alumina. The total linear superficial gas velocity in the fluidized bed was about 1 to 2 ft./ sec. The feed was introduced into the first stage reactor as a gas containing entrained liquid hydrocarbons and solid catalyst, where it was heated to about 850 F. at a pressure of about 8 p.s.i.g. and a WHSV of about 15. Under these conditions, a 30-40% conversion of the feed to lighter materials was effected with the eflluent being substantially free of the metal contaminants along with the associated coke formers. Catalyst is taken from the reactor and its carbon content is reduced from about 2 to 0.5 weight percent through contact with air in the regenerator.

A slip stream of regenerated catalyst analyzing 0.5% coke, 550 p.p.m. nickel and 1570 p.p.m. vanadium is continuously withdrawn from the regenerator of the first stage cracker at a daily inventory rate of 60% and sent to demetallization where it is held for about 2 hours in a zone where it is contacted with air at about 1300 F. and then sent to a sulfiding zone where it is fluidized with H S gas at a temperature of about 1050 F. for about 2 hours. Water containing dilute hydrogen peroxide mixed with nitric acid is brought in contact with the sulfided catalyst for minutes at a temperature of 200 F. The catalyst is then washed with an ammonium hydroxide solution having a pH of about 8 to 11, removing the available vanadium. The metals level of the cracking catalyst after demetallization analyzes 275 p.p.m. nickel and 1150 p.p.m. vanadium.

The cracked products from the first cracking zone are introduced into a fractionator where the products were separated into a gasoline fraction having an end boiling point of 430 F. which is recovered, and a cycle oil residual fraction boiling between about 400 F. to 850- 950 F. This 400 F. plus fraction having a metals content of about 0.20 p.p.m. nickel and 0.4 p.p.m. vanadium is directed to a second catalytic cracking zone where this cycle oil is cracked over a clean and selective silicaalumina catalyst of the composition of the first stage catalyst, at a temperature of about 950 F. and a pressure of about 10-20 p.s.i.g. with a WHSV of about 5 to give a conversion of 75-80% by volume of the cycle oil introduced into the second cracking zone. A portion of the silica-alumina catalyst is removed from the second stage reactor where it is introduced into a regenerator and kept in a free oxygen-containing gas at about 1050 F. for about 9 minutes.

It is apparent from the foregoing that it is now possible to provide a process for the conversion in a high yield of residua containing metallic impurities which have harmful efiects on the selectivity of a cracking catalyst, by treating the residuum in a low severity first stage cracking unit having associated with this unit demetallization facilities to keep the metallic content of the catalyst below harmful levels, followed by extensive conversion of the efliuent from the first stage over a relatively non-contaminated, selective catalyst in a second cracking stage.

It is claimed:

1. A process for catalytically cracking a hydrocarbon feedstock to produce gasoline, said feedstock consisting essentially of residual petroleum hydrocarbon boiling above about 600 F. and containing more than about 0.6 part per million nickel and more than 1.5 parts per million vanadium, which comprises subjecting said feedstock to first-stage catalytic cracking employing a synthetic gel, silica-based cracking catalyst at a temperature in the range of about 700 to 900 F., and a pressure of about 5 to 15 p.s.i.g., thereby depositing substantially all of said metallic impurities on said catalyst and converting about 30 to 45% of said petroleum stock to lower boiling materials, withdrawing from said first stage cracking system contaminated catalyst containing at least about parts per million nickel and at least about 250 parts per million vanadium, demetallizing the withdrawn catalyst and returning the demetallized catalyst to said first stage catalytic cracking system, fractionating the products from the first stage to separate a bottoms fraction boiling essentially above 400 F. containing no more than about 0.4 part per million nickel and no more than 0.8 part per million vanadium and a gasoline fraction, and subjecting the said bottoms fraction to second stage catalytic cracking employing a silica-based cracking catalyst at a temperature at least about 25 F. higher and at a higher cracking severity than used in said first stage.

.2. The process of claim ll, wherein the temperature of the first cracking stage is in the range of about 800 to 875 F.

3. The process of claim 2 wherein the temperature of the second cracking stage is about 900 to 1000 F. and the pressure is about 5 to 25 p.s.i.g.

4. The process of claim 3 wherein the catalyst of the first cracking zone is a synthetic gel silica-alumina catalyst of low activity.

5. The process of claim 1 wherein catalyst demetallization includes contact at an elevated temperature with vapors reactive with the poisoning metal.

6. The process of claim 1 wherein the first stage catalytic cracking is conducted in the presence of a diluent selected from the group consisting essentially of inert gases and low end point hydrocarbon vapors.

7. The process of claim 6 wherein the diluent is steam in amounts ranging from about 20 to 40 weight percent based on the amount of residual petroleum oil feed.

8. A process for catalytically cracking a hydrocarbon feedstock to produce gasoline, said feedstock consisting essentially of residual petroleum hydrocarbon boiling above about 600 F. and containing more than about 0.6 part per million nickel and more than 1.5 parts per million vanadium, which comprises subjecting said feedstock to first-stage catalytic cracking employing a synthetic gel, silica-based cracking catalyst at a temperature in the range of about 700 to 900 F., a pressure of about 5 to 15 p.s.i.g. and a WHSV of about 10 to 20, thereby depositing substantially all of said metallic impurities on said catalyst and converting about 30 to 45% of said petroleum stock to lower boiling materials, withdrawing from said first stage cracking system contaminated catalyst containing at least about 100 parts per million nickel and at least about 250 parts per million vanadium, contacting the catalyst after regeneration to remove carbon, with a molecular oxygen-containing gas at a temperature of about 1000 to about 1800 F., sulfiding the poisoning metal containing component on the catalyst by contact with a sulfiding agent at a temperature of about 500 to 1500 F., chlorinating poisoning-metal containing component on the catalyst by contact with an essentially anhydrous chlorinating agent at a temperature of about 300 to 1000 F., contacting the catalyst with a liquid essentially aqueous medium to remove soluble poisoning metal chloride from the catalyst and returning the demetallized catalyst to said first stage catalytic cracking system, fractionating the roducts from the first stage to separate a bottoms fraction boiling essentially above 400 F. containing no more than about 0.4 part per million nickel and no more than about 0.8 part per million vanadium and a gasoline fraction, and subjecting the bottoms fraction to second stage catalytic cracking employing a silica-based cracking catalyst at a temperature of at least about 25 F. higher and at a higher cracking severity than used in said first stage.

9. The process of claim 8 in which the sulfiding is performed by contact with H 5.

30. The process of claim 8 in which the chlorinating is performed with an equimolar mixture of C1 and CCl References (Iited by the Examiner UNITED STATES PATENTS Gray 208-113 Shabaker et al. 208-120 Snyder 208120 Snyder 20S113 Blending 20874 Mattox et a1 208-251 Conn et a1. 208251 ALPHONSO D. SULLIVAN, Primary Examiner.

MELTON STERMAN, Examiner. 

1. A PROCESS FOR CATALYTICALLY CRACKING A HYDROCARBON FEEDSTOCK TO PRODUCE GASOLINE, SAID FEEDSTOCK CONSISTING ESSENTIALLY OF RESIDUAL PETROLEUM HYDROCARBON BOILING ABOVE ABOUT 600*F. AND CONTAINING MORE THAN ABOUT 0.6 PART PER MILLION NICKEL AND MORE THAN 1.5 PARTS PER MILLION VANADIUM, WHICH COMPRISES SUBJECTING SAID FEEDSTOCK TO FIRST-STAGE CATALYTIC CRACKING EMPLOYING A SYNTHETIC GEL, SILICA-BASED CRACKING CATALYST AT A TEMPERATURE IN THE RANGE OF ABOUT 700 TO 900*F., AND A PRESSURE OF ABOUT 5 TO 15 P.S.I.G., THEREBY DEPOSITING SUBSTANTIALLY ALL OF SAID METALLIC IMPURITIES ON SAID CATALYST AND CONVERTING ABOUT 30 TO 45% OF SAID PETROLEUM STOCK TO LOWER BOILING MATERIALS, WITHDRAWING FROM SIAD FIRST STAGE CRACKING SYSTEM CONTAMINATED CATALYST CONTAINING AT LEAST ABOUT 100 PARTS 